CA2004882A1 - Process for reducing coke formation during hydroconversion of heavy hydrocarbons - Google Patents
Process for reducing coke formation during hydroconversion of heavy hydrocarbonsInfo
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- CA2004882A1 CA2004882A1 CA 2004882 CA2004882A CA2004882A1 CA 2004882 A1 CA2004882 A1 CA 2004882A1 CA 2004882 CA2004882 CA 2004882 CA 2004882 A CA2004882 A CA 2004882A CA 2004882 A1 CA2004882 A1 CA 2004882A1
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- reactor
- diluent
- reaction mixture
- feedstock
- hydrogen gas
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Abstract
"PROCESS FOR REDUCING COKE FORMATION DURING
HYDROCONVERSION OF HEAVY HYDROCARBONS"
Abstract of the Disclosure A process is provided for hydroconverting a heavy hydrocarbon feedstock. A diluent is admixed with the feedstock and the reaction mixture is preheated to a temperature of about 350°C. Optionally, a metal carbonyl compound may be added to the reaction mixture. The reaction mixture is contacted with hydrogen gas in a reactor having hydrogen partial pressure. The hydrogen gas flow rate is adapted to substantially immediately strip off the generated lighter volatile components. Thus the region of separation of the liquid phase into two or more separate liquid phases may be avoided.
HYDROCONVERSION OF HEAVY HYDROCARBONS"
Abstract of the Disclosure A process is provided for hydroconverting a heavy hydrocarbon feedstock. A diluent is admixed with the feedstock and the reaction mixture is preheated to a temperature of about 350°C. Optionally, a metal carbonyl compound may be added to the reaction mixture. The reaction mixture is contacted with hydrogen gas in a reactor having hydrogen partial pressure. The hydrogen gas flow rate is adapted to substantially immediately strip off the generated lighter volatile components. Thus the region of separation of the liquid phase into two or more separate liquid phases may be avoided.
Description
~:0048a2 ~
2 The present invention relates to an improvement in a 3 hydroconversion process for the upgrading of heavy hydrocarbon 4 feedstocks. ~ .
"
BACKGRO~ND OF TEE INVENTION
6 The processes which are used to upgrade bitumen are 7 thermal cracking, hydrocracking and hydrotreating.
8 It has long been recognized that the formation of a g separate phase, the mesophase, in bitumen or residua that is being hydroconverted presages the formation of coke.
11 Deleteriously, this coke may adhere to surfaces or when suspended 12 in process fluid~ agglomerate and hinder flow. Microscopic 13 methods may be applied to detect the separation of the liquid 14 into two or more separate phases for the subsequent onset of coke formation.
16 By way of background it is known that at ambient 17 conditions, the addition of an aliquot of pentane to bitumen will 18 cause the asphaltenes to precipitate out as either a separate 19 liquid or solid phase depending upon the quantity of pentane added. In a similar manner, it is believed that in a thermal 21 cracker, asphaltenes and other large molecules separate from the 22 reacting mixture as hydrogen dissolves and light paraffins are 23 formed in the mixture. Stated otherwise, beyond a certain ! " ~, 24 conversion level, the composition of the reactor fluid becomes unstable. This is due to the cumulative effects of 26 incompatibility of solution properties between the light volatile 27 products and the heavy hydrocarbons. Consequently, phase 28 separatlon of the heavy hydrocarcons wlll take place. Subsequent ~ .
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1 cooling of the mixture will cause the precipitates to settle ou~
2 and deposit as carbonaceous solids in the product lines.
3 Prolonged residence of the separated phase in the reactor rapidly 4 results in its conversion to coke. Either occurrence will limit the feed conversion to the desired light products and result in 6 e~cessive coke formation and catalyst deactivation.
7 In Canadian Patent 653,618, there is disclosed a two-8 stage hydrogenation process which attempts to address the g inherent incompatibility between the lighter fraction and the heavier material. During the first process stage, the feed oil 11 is passed downwardly through a thermal hydrogenation~
12 fractionation zone. Hydrogen is passed counter currently to the 13 feed oil via bubble cap plates. One of the major problems 14 associated with this process concerns the fouling of both the reactor column and the bubble cap plates due to the adhesion of 16 coke thereon.
17 It was disclosed in U.S. Patent 4,411,768, that the 18 coke precursors could be separated from the recycle stream of an 19 ebullated catalytic bed reactor. Nevertheless, coking and solids deposition continues to take place during the practise of this 21 process.
22 U.S. Patent 4, 443,328, discloses a thermal process 23 wherein a preheated heavy oil feedstock is charged into the upper 24 reaction zone of an upright, cylindrical continuous reactor. The reactor is divided into a plurality of discreet reaction zones by 26 the provision of partition plates. Steam is passed upwardly, 27 countercurrent to the downwardly flowing heavy liquid. Gas and 28 oil vapour are removed from the reactor top. The residual pitch 29 is discharged from the bottom of the reactor. A rotary scraper 20~48~2 1 prevents excessive coke deposition. Notwithstanding, excessive 2 coking problems limit the conversion to about 60 3 A process in which heavy hydrocarbon oil is 4 hydrocracked in the presence of a hydrogen donor is described in U.S. Patent 4,485,004. A particulate hydrogenation catalyst is 6 used.
7 A two-stage petroleum residuum hydroconversion is 8 described in Canadian Patent 1,242,403. The first stage uses g counterflow gas-liquid hydrothermal conditions. In the second process stage, the heavy hydrocarbons produced from the first 11 stage, are hydroprocessed at a much lower temperature over a 12 fixed bed comprising a hydrocracking catalyst.
13 At the time of the present invention, therefore, there 14 still existed the need for a process functional to overcome the constraints imposed on conversion limits. These constraints 16 arose from the insolubility of the light products in the heavy 17 hydrocarbons.
19 The present invention is founded on the recognition that it is possible to control the phase separation during the 21 hydroconversion process, avoiding the region of separation of the 22 liquid phase into two or more separate liquid phases. This is 23 attained by the simple expedient of providing a high gaseous flow 24 rate through the reactor and, preferably, by addition of a diluent to the feedstock. By controlling phase separation is 26 meant in accordance with predictions using a simple ternary phase 27 diagram. The preferred gas would be hydrogen. Most preferably, 28 a recycle hydrogen gas stream would be used. However, one could ~:0048~
1 contemplate the use of alternative suitable gases, such as 2 nitrogen or the like. This high gas flow rate, preferably in 3 combination with the added diluent, functions to immediately 4 segregate the light and heavy products in the reactor thereby continuously creating a compatible reactor fluid therein for the 6 heavy fraction to react without precipitation. Stated otherwise, 7 one seeks to maximize stripping off of the lighter products upon 8 their generation, during each stage of the hydroconversion 9 reactions to thereby inhibit the deasphalting process.
This discovery was based on the observation that, if 11 when the feed is about 50% converted, the light products are 12 permitted to remain in the reactor, then asphaltenes and other 13 heavy hydrocarbon molecules will precipitate with concomitant 14 generation of additional coke.
Preferably, the hydrocracking of the less crackable 16 heavy molecules i8 carried out at a high hydrogen overpressure in 17 the reactor.
18 It has also been found preferable to combine the heavy 19 hydrocarbon feedstock with a diluent. An exemplary diluent could comprise a product stream generated by cracking heavy oils, said 21 diluent having a boiling point between 200C and 504C.
22 Furthermore, a homogeneous additive capable of 23 minimizing coke forming reactions may be used. The additive 24 preferably is a homogeneous metal complex such as an iron, nickel, cobalt or molybdenum carbonyl.
26 As a result of practising the process it has been 27 possible to obtain extremely high conversion rates of the order 28 of 98% with concomitantly much reduced coke formation. As an 29 added advantage, the process may be conducted without the :~
20~48 1 presence of a catalyst. ~ ~
2 In a broad aspect, the invention is a process for ~ ~.
3 hydroconverting a heavy hydrocarbon feedstock which comprises:
4 - admixing a diluent with the feedstock and preheating the reaction mixture to a temperature ~ ~;
ç of at least about 350C; and 7 - contacting the reaction mixture with hydrogen gas 8 in a reactor having hydrogen partial pressure, 9 said hydrogen gas having a hydrogen concentration greater than 70% and having a flow rate of between 11 about 2,000 - 18,000 scf/BBL, at a liquid hourly 12 space velocity of between about 0.2 to about 2.0, 13 whereby the hydrogen gas flow rate is adapted to 14 substantially immediately strip off the generated lighter volatile components at a reaction 16 temperature in the range of between about 420C to 17 about 480C to thereby selectively hydroconvert -18 said heavy hydrocarbon feedstock and minimize ~ ~5 19 carbonaceous precipitation therein.
DESCRIPTION OF THE DRAWINGS
21 Figure 1 is a preferred embodiment of a flowsheet which 22 could be utilized in the practise of the process of the present ~ i -23 invention.
24 Figure 2 is an alternative flowsheet usable in the process of the instant invention.
26 Figure 3 illustrates the effect of the continuous 27 removal of the light volatile components on the coke formation. ~
28 Figure 4 is a hypothetical ternary diagram for -6 `s .: :, ` ,' .'~;:'''.:
~' ~"''' 20~48~32 1 pseudocomponents of bitumen. The arrow shows the path of 2 changing composition that might occur as the bitumen is heated 3 and hydrocracking occurs.
Having reference to the accompanying drawings, there is 6 provided an improvement to a hydroconversion process which is 7 applicable to heavy hydrocarbons.
8 Whilst the present embodiment is being described having g particular reference to a process involving the addition of a donor diluent it will be readily appreciated that the principles 11 underlying the present process may be applied to similar ~
12 processes by one ~killed in the art. ~i;
13 The feedstock and, optionally, a diluent are introduced 14 into the reactor 10 via a line 2.
The reaction mixture is preheated to a temperature of 16 at least about 350C in a conventional preheater (not shown).
17 The feedstock may be selected from a petroleum resid, 18 coal-oil coprocessing liquid or coal liquification liquid.
19 Typically however, the feedstock would comprise a bitumen or heavy oil or vacuum-tower-bottom residuum having an initial 21 boiling point (IBP) of approximately 504C.
22 The donor diluent preferably comprises a product stream 23 generated by cracking heavy oils. The diluent would preferably 24 have a boiling point between 220C and 450C. The concentration of feedstock to diluent may range from 10:1 to about 1:1.
26 A homogeneous additive may optionally be introduced 27 with the feedstock. The additive may comprise iron, cobalt, 28 nickel or molybdenum carbonyl. The metal concentration would ;~0048~.
1 range from about 10 to about 5,000 p.p.m. Alternatively, a metal 2 salt such as iron sulphate may be used.
3 The preferred reactor de~ign is that similar to a 4 continuous stirred tank reactor (CSTR), having a height to diameter ratio of the order of 5. However, the height of free 6 flow within the reactor under high gas flow is a concern since 7 the upward flowing bubbles of gas break up to form a foam.
8 Simple remedies such as providing sequential reactors, or g providing baffles or other means of disturbing the fluid flow to accomplish the same, or adding antifoaming agents such as 11 silicone based oils may be required.
12 Hydrogen gas is injected into the base of the reactor 13 via a line 3. The hydrôgen gas is preferably recycle gas from 14 downstream, although some make-up gas will be required. The hydrogen gas would require a flow rate of between about 2,000 to 16 18,000 scf/BBL at a liquid hourly space velocity (LHSV) of 17 between about 0.2 to 2Ø Preferably the LHSV would range from 18 about 0.5 to 1Ø The reaction temperature would range from 19 between about 420C to about 480C.
As will be evident to one skilled in the art, the 21 reaction is dependent on the parametric properties of hydrogen 22 flow rate, reaction temperature, choice of diluent, additive 23 properties and the degree of conversion of the feedstock. Thus 24 under hydrocracking conditions the hydrogen gas flow rate would range from between about 3,000 to about 18,000 scf/BBL at a 26 liquid hourly space velocity of between 0.2 to 2.0 at a reaction 27 temperature ranging between 455C to 480C. Under hydrotreating 28 conditions, the hydrogen gas flow rate would range between about 29 2,000 scftBBL to about 5,000 scf/BBL at a liquid hourly space ; :0~48~
1 velocity of between about 0.5 to about 2.0 at a reaction 2 temperature in the range of between about 420C to about 455C.
3 The preferred hydrogen overpressure in the reactor 4 would range from 3.5 MPa to about 15 MPa. Preferably, the hydrogen overpressure would range from about 7 NPa to about 15 6 MPa.
7 In a CSTR operating mode, the hydrogen gas flow could 8 be cocurrent. However, it is preferred that the hydrogen gas be g sparged into the reactor in countercurrent flow with said feedstock. ;;
11 The overhead stream from the reactor 10 is passed via a ~ -12 line 4 to a conventional separator 5 where the VGO is withdrawn 13 from the base. The overhead stream is passed via a line 7 14 through a condenser 8 and into a second separator 9. The bottom stream comprising naphtha and the higher naphtha fractions are 16 withdrawn from the base of the separator 9. The overhead stream ~;
17 comprising hydrogen and the light gases after further scrubbing 18 are passed into a pressure swing adsorption unit 11. The recycle 19 hydrogen stream from the unit 11 is passed to the reactor via a line 12. Make-up hydrogen gas may be introduced into the line 12 21 via line 13.
22 It is to be noted that the present process may be 23 practised using a pair of reactors 10 and 14 respectively, in 24 series as illustrated in Figure 2.
The feedstock and diluent are introduced into the first 26 reactor 10 via a line 15. A hydrogen stream is also sparged into 27 the base of the reactor 10. The overheads are passed via a line 28 16 into a first separator 17. Hydrogen together with a diluent 29 is introduced lnto the ~eparator 17. The fluid product~ from the . " ,;~
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1 separator 17 are pa~sed via a line 18 into the second reactor 14.
2 Hydrogen is sparged into the base of the reactor 14 via a line 3 18. The overhead stream from the reactor 14 is further proce~ed 4 conventionally. Alternatively, the overhead stream may be passed to a further reactor in series and treated as described supra.
6 The operating condition such as temperature and gas flow rates 7 would be as described hereinabove.
g The following experimental results are included to describe the operability of the instant invention.
11 A one litre autoclave was fitted with a mechanical 12 stirrer and baffles, a dip tube for sparging N2 or H2 into the 13 reacting liquid, an outlet permitting continuous flow of product 14 gas, and cold trap condensers to remove volatile products from the gas stream before collecting the latter in a sample bag for 16 analysis. Experiments were conducted in the above described 17 reactor at 430C for 105 minutes under 550 RPa pressure both 18 without continuou4 gas flow and with continuous gas flowing in 19 and out of the reactor. In each experiment, 110 g of Athabasca vacuum tower bottom (VTB) and 220 g of a diluent were used.
21 Table I presented herebelow gives the detailed reaction 22 conditions and experimental results.
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- ~0048~32 2Batch Autoclave Test Results 3 Test No~ 1 2 3 4 5 6 :
4 Diluent TypeA A B B B B
Nitrogen Flow :
6 Rate (l/min)0.01.89 0-0 0.2 1.05 2.16 -~
7 Yield (wt% VTB):
8 C1-C4 15.5 16.510.3 10.0 8.2 13.8 :~
g Cs - 504C- 42.3 38.156.5 56.~ 56.6 51.4 504C+ pitch 11 (coke free)34.638.826.3 27.0 30.0 29.7 12 Coke 7.9 3.8 4.3 3.8 3.4 3.1 13 Condensate 14 recovered from :.
purge gas (wt% VTB) --47.5 -- 3.1 23.9 39.7 17 Simulated Distillation Results of Condensate from Test 5 18% Off Temp C % Off Temp C
19 IBP 34 55 146 .
24 25 98 80 201 :~
111 85 210 ~-27 40 123 95 229 ~-`
28 45 131 FBP 262 ~ .
29 50 139 -- --- ~
11 ' ' ';
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Table III herebelow provides typical diluent and ~-2 Athabasca Vacuum Tower bottom compositions.
4 Diluent Type A B ~ -Weight %~
6 Paraffins 13.0 16.4 ~;
7 Uncondensed 8 Cycloparaffins 7.2 6.3 g Condensed ~ .
Dicycloparaffins 5.2 13.1 ~ :
11 Alkylbenzenes 18.1 15.3 12 Benzocycloparaffins32.3 37.5 :~
13 Benzodicycloparaffin~ 4.8 3.8 ~ ~-14 Naphthalenes 15.9 6.1 . ... -.
Athabasca Vacuum Tower Bottoms (Wt%) ~ -16 C - 81.76 H - 9.51 S -- 6.23 N - 0.78 - -17 API @ 16C : 2.43 ~ ;
"~
-. :'' 18 The large differences in product yield between tests 19 conducted with the respective diluents A and B are due to the high concentration of substituted naphthalenes and tetralins in 21 the latter.
22 In addition to the above described tests, continuous ~ ~
23 flow experiments were performed in a single stirred tank reactor .`
24 and two stirred tank reactors connected in series. A mixture of .:
VTB, diluent, additive (optionally used), and preheated hydrogen ;.
26 were fed into a preheater coil immersed in a heated fluidized l~, '''~' .
20048~
1 sandbath. Hydrocracking reactions were conducted over the 2 appropriate r~sidence time, measured as Liquid Hourly Space ~:
3 Velocity, in the reactor arrangement described in Table IV
4 herebelow. A hot separator split the product into a gas plus volatile stream and a liquid slurry stream.
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;:0048~32 2 Continuous Bench Unit Test Results ;::
3 Diluent Type : A ~
4 Pressure : 10 NPa ~;
' ',.~, S Test No. 7 8 9 10 6 Reactor 1 (1) 1.2 1.2 1.2 0.5 7 Reactor 2 (1) -- -- -- 0.7 ;~;
8 Reaction Temperature 9 (C) 440 440 460 440 Liquid Hourly ~-11 Space Velocity (hr~l) 1 0.5 0.5 12 Hydrogen Flow rate 13 (slpm) 12 12 12 16 14 VTB Concentration in feed (Wt%~63.10 65.0245.5347.3 16 Conversion (Wt% VTB
17 to coke and 504C-)69.0 78.398.1 79.8 ', 18 Yield, Wt% VTB
19 C1 - C4 5.2 7.9 17.3 7.2 C5 - 200C 16.3 19.1 31.4 15.1 21 200C - 360C 23.4 31.1 43.1 24. 2 : ~:
22 360C - 504C 20. 3 18.4 9. 3 24.8 ~ -~
23 Coke 4.6 3.1 0.1 4.6 24 504C+ Pitch (coke free) 31. 0 21.7 1.9 24.8 ,~
26 Total distillate 27 Cs - 504C 60.0 68.683. 8 64.7 28 It is to be noted that in the test run 9 it was 29 interrupted shortly after the yield period. Plugging occurred downstream of the reactor. Nevertheless, the reactor remained 31 clean in spite of shutdown.
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2Table V herebelow shows data with an additive.
TABLE v --~
3 Run No. 11 12 4 Reactor l (1) 1.2 1.2 Reactor 2 (1) 6 Reaction 7 Temperature (C) 450 450 8 Liquid Hourly Space g Velocity (hr~1) 1.05 0.73 (slpm) 8 12 12 VTB Concentration in Feed 13 (wt%) 47.5 47.8 14 Catalyst (Nt~ of Fe based on VTB) 0.5 0.5 ;~
16 Conversion (Wt% VTB
17 to coke and 504C-) 67.8 82.1 18 Yield (Wt% VTB) 19 Cl-C4 5.2 9.4 Cs-200C 19.1 22.9 21 200-350C l9.9 27.6 22 350-504C 21.9 21.1 23 Coke 2.4 2.8 24 504C+ Pitch Coke free 32.2 17.9 ..
26 Total distillate 27 Cs-504C 60.9 71.6 . -~ . .
28 Note that if H2 flow wa~ not increased in the test run 12, one 29 would expect that a 14% increase in conversion should be .::
accompanied by much higher coke yield than the 0.4~ recorded.
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"
BACKGRO~ND OF TEE INVENTION
6 The processes which are used to upgrade bitumen are 7 thermal cracking, hydrocracking and hydrotreating.
8 It has long been recognized that the formation of a g separate phase, the mesophase, in bitumen or residua that is being hydroconverted presages the formation of coke.
11 Deleteriously, this coke may adhere to surfaces or when suspended 12 in process fluid~ agglomerate and hinder flow. Microscopic 13 methods may be applied to detect the separation of the liquid 14 into two or more separate phases for the subsequent onset of coke formation.
16 By way of background it is known that at ambient 17 conditions, the addition of an aliquot of pentane to bitumen will 18 cause the asphaltenes to precipitate out as either a separate 19 liquid or solid phase depending upon the quantity of pentane added. In a similar manner, it is believed that in a thermal 21 cracker, asphaltenes and other large molecules separate from the 22 reacting mixture as hydrogen dissolves and light paraffins are 23 formed in the mixture. Stated otherwise, beyond a certain ! " ~, 24 conversion level, the composition of the reactor fluid becomes unstable. This is due to the cumulative effects of 26 incompatibility of solution properties between the light volatile 27 products and the heavy hydrocarbons. Consequently, phase 28 separatlon of the heavy hydrocarcons wlll take place. Subsequent ~ .
`'.:,,.
,~
~''`'."
20~48~Z ~;
1 cooling of the mixture will cause the precipitates to settle ou~
2 and deposit as carbonaceous solids in the product lines.
3 Prolonged residence of the separated phase in the reactor rapidly 4 results in its conversion to coke. Either occurrence will limit the feed conversion to the desired light products and result in 6 e~cessive coke formation and catalyst deactivation.
7 In Canadian Patent 653,618, there is disclosed a two-8 stage hydrogenation process which attempts to address the g inherent incompatibility between the lighter fraction and the heavier material. During the first process stage, the feed oil 11 is passed downwardly through a thermal hydrogenation~
12 fractionation zone. Hydrogen is passed counter currently to the 13 feed oil via bubble cap plates. One of the major problems 14 associated with this process concerns the fouling of both the reactor column and the bubble cap plates due to the adhesion of 16 coke thereon.
17 It was disclosed in U.S. Patent 4,411,768, that the 18 coke precursors could be separated from the recycle stream of an 19 ebullated catalytic bed reactor. Nevertheless, coking and solids deposition continues to take place during the practise of this 21 process.
22 U.S. Patent 4, 443,328, discloses a thermal process 23 wherein a preheated heavy oil feedstock is charged into the upper 24 reaction zone of an upright, cylindrical continuous reactor. The reactor is divided into a plurality of discreet reaction zones by 26 the provision of partition plates. Steam is passed upwardly, 27 countercurrent to the downwardly flowing heavy liquid. Gas and 28 oil vapour are removed from the reactor top. The residual pitch 29 is discharged from the bottom of the reactor. A rotary scraper 20~48~2 1 prevents excessive coke deposition. Notwithstanding, excessive 2 coking problems limit the conversion to about 60 3 A process in which heavy hydrocarbon oil is 4 hydrocracked in the presence of a hydrogen donor is described in U.S. Patent 4,485,004. A particulate hydrogenation catalyst is 6 used.
7 A two-stage petroleum residuum hydroconversion is 8 described in Canadian Patent 1,242,403. The first stage uses g counterflow gas-liquid hydrothermal conditions. In the second process stage, the heavy hydrocarbons produced from the first 11 stage, are hydroprocessed at a much lower temperature over a 12 fixed bed comprising a hydrocracking catalyst.
13 At the time of the present invention, therefore, there 14 still existed the need for a process functional to overcome the constraints imposed on conversion limits. These constraints 16 arose from the insolubility of the light products in the heavy 17 hydrocarbons.
19 The present invention is founded on the recognition that it is possible to control the phase separation during the 21 hydroconversion process, avoiding the region of separation of the 22 liquid phase into two or more separate liquid phases. This is 23 attained by the simple expedient of providing a high gaseous flow 24 rate through the reactor and, preferably, by addition of a diluent to the feedstock. By controlling phase separation is 26 meant in accordance with predictions using a simple ternary phase 27 diagram. The preferred gas would be hydrogen. Most preferably, 28 a recycle hydrogen gas stream would be used. However, one could ~:0048~
1 contemplate the use of alternative suitable gases, such as 2 nitrogen or the like. This high gas flow rate, preferably in 3 combination with the added diluent, functions to immediately 4 segregate the light and heavy products in the reactor thereby continuously creating a compatible reactor fluid therein for the 6 heavy fraction to react without precipitation. Stated otherwise, 7 one seeks to maximize stripping off of the lighter products upon 8 their generation, during each stage of the hydroconversion 9 reactions to thereby inhibit the deasphalting process.
This discovery was based on the observation that, if 11 when the feed is about 50% converted, the light products are 12 permitted to remain in the reactor, then asphaltenes and other 13 heavy hydrocarbon molecules will precipitate with concomitant 14 generation of additional coke.
Preferably, the hydrocracking of the less crackable 16 heavy molecules i8 carried out at a high hydrogen overpressure in 17 the reactor.
18 It has also been found preferable to combine the heavy 19 hydrocarbon feedstock with a diluent. An exemplary diluent could comprise a product stream generated by cracking heavy oils, said 21 diluent having a boiling point between 200C and 504C.
22 Furthermore, a homogeneous additive capable of 23 minimizing coke forming reactions may be used. The additive 24 preferably is a homogeneous metal complex such as an iron, nickel, cobalt or molybdenum carbonyl.
26 As a result of practising the process it has been 27 possible to obtain extremely high conversion rates of the order 28 of 98% with concomitantly much reduced coke formation. As an 29 added advantage, the process may be conducted without the :~
20~48 1 presence of a catalyst. ~ ~
2 In a broad aspect, the invention is a process for ~ ~.
3 hydroconverting a heavy hydrocarbon feedstock which comprises:
4 - admixing a diluent with the feedstock and preheating the reaction mixture to a temperature ~ ~;
ç of at least about 350C; and 7 - contacting the reaction mixture with hydrogen gas 8 in a reactor having hydrogen partial pressure, 9 said hydrogen gas having a hydrogen concentration greater than 70% and having a flow rate of between 11 about 2,000 - 18,000 scf/BBL, at a liquid hourly 12 space velocity of between about 0.2 to about 2.0, 13 whereby the hydrogen gas flow rate is adapted to 14 substantially immediately strip off the generated lighter volatile components at a reaction 16 temperature in the range of between about 420C to 17 about 480C to thereby selectively hydroconvert -18 said heavy hydrocarbon feedstock and minimize ~ ~5 19 carbonaceous precipitation therein.
DESCRIPTION OF THE DRAWINGS
21 Figure 1 is a preferred embodiment of a flowsheet which 22 could be utilized in the practise of the process of the present ~ i -23 invention.
24 Figure 2 is an alternative flowsheet usable in the process of the instant invention.
26 Figure 3 illustrates the effect of the continuous 27 removal of the light volatile components on the coke formation. ~
28 Figure 4 is a hypothetical ternary diagram for -6 `s .: :, ` ,' .'~;:'''.:
~' ~"''' 20~48~32 1 pseudocomponents of bitumen. The arrow shows the path of 2 changing composition that might occur as the bitumen is heated 3 and hydrocracking occurs.
Having reference to the accompanying drawings, there is 6 provided an improvement to a hydroconversion process which is 7 applicable to heavy hydrocarbons.
8 Whilst the present embodiment is being described having g particular reference to a process involving the addition of a donor diluent it will be readily appreciated that the principles 11 underlying the present process may be applied to similar ~
12 processes by one ~killed in the art. ~i;
13 The feedstock and, optionally, a diluent are introduced 14 into the reactor 10 via a line 2.
The reaction mixture is preheated to a temperature of 16 at least about 350C in a conventional preheater (not shown).
17 The feedstock may be selected from a petroleum resid, 18 coal-oil coprocessing liquid or coal liquification liquid.
19 Typically however, the feedstock would comprise a bitumen or heavy oil or vacuum-tower-bottom residuum having an initial 21 boiling point (IBP) of approximately 504C.
22 The donor diluent preferably comprises a product stream 23 generated by cracking heavy oils. The diluent would preferably 24 have a boiling point between 220C and 450C. The concentration of feedstock to diluent may range from 10:1 to about 1:1.
26 A homogeneous additive may optionally be introduced 27 with the feedstock. The additive may comprise iron, cobalt, 28 nickel or molybdenum carbonyl. The metal concentration would ;~0048~.
1 range from about 10 to about 5,000 p.p.m. Alternatively, a metal 2 salt such as iron sulphate may be used.
3 The preferred reactor de~ign is that similar to a 4 continuous stirred tank reactor (CSTR), having a height to diameter ratio of the order of 5. However, the height of free 6 flow within the reactor under high gas flow is a concern since 7 the upward flowing bubbles of gas break up to form a foam.
8 Simple remedies such as providing sequential reactors, or g providing baffles or other means of disturbing the fluid flow to accomplish the same, or adding antifoaming agents such as 11 silicone based oils may be required.
12 Hydrogen gas is injected into the base of the reactor 13 via a line 3. The hydrôgen gas is preferably recycle gas from 14 downstream, although some make-up gas will be required. The hydrogen gas would require a flow rate of between about 2,000 to 16 18,000 scf/BBL at a liquid hourly space velocity (LHSV) of 17 between about 0.2 to 2Ø Preferably the LHSV would range from 18 about 0.5 to 1Ø The reaction temperature would range from 19 between about 420C to about 480C.
As will be evident to one skilled in the art, the 21 reaction is dependent on the parametric properties of hydrogen 22 flow rate, reaction temperature, choice of diluent, additive 23 properties and the degree of conversion of the feedstock. Thus 24 under hydrocracking conditions the hydrogen gas flow rate would range from between about 3,000 to about 18,000 scf/BBL at a 26 liquid hourly space velocity of between 0.2 to 2.0 at a reaction 27 temperature ranging between 455C to 480C. Under hydrotreating 28 conditions, the hydrogen gas flow rate would range between about 29 2,000 scftBBL to about 5,000 scf/BBL at a liquid hourly space ; :0~48~
1 velocity of between about 0.5 to about 2.0 at a reaction 2 temperature in the range of between about 420C to about 455C.
3 The preferred hydrogen overpressure in the reactor 4 would range from 3.5 MPa to about 15 MPa. Preferably, the hydrogen overpressure would range from about 7 NPa to about 15 6 MPa.
7 In a CSTR operating mode, the hydrogen gas flow could 8 be cocurrent. However, it is preferred that the hydrogen gas be g sparged into the reactor in countercurrent flow with said feedstock. ;;
11 The overhead stream from the reactor 10 is passed via a ~ -12 line 4 to a conventional separator 5 where the VGO is withdrawn 13 from the base. The overhead stream is passed via a line 7 14 through a condenser 8 and into a second separator 9. The bottom stream comprising naphtha and the higher naphtha fractions are 16 withdrawn from the base of the separator 9. The overhead stream ~;
17 comprising hydrogen and the light gases after further scrubbing 18 are passed into a pressure swing adsorption unit 11. The recycle 19 hydrogen stream from the unit 11 is passed to the reactor via a line 12. Make-up hydrogen gas may be introduced into the line 12 21 via line 13.
22 It is to be noted that the present process may be 23 practised using a pair of reactors 10 and 14 respectively, in 24 series as illustrated in Figure 2.
The feedstock and diluent are introduced into the first 26 reactor 10 via a line 15. A hydrogen stream is also sparged into 27 the base of the reactor 10. The overheads are passed via a line 28 16 into a first separator 17. Hydrogen together with a diluent 29 is introduced lnto the ~eparator 17. The fluid product~ from the . " ,;~
:~,:- .
-:' '` , 20C~4~
: .
1 separator 17 are pa~sed via a line 18 into the second reactor 14.
2 Hydrogen is sparged into the base of the reactor 14 via a line 3 18. The overhead stream from the reactor 14 is further proce~ed 4 conventionally. Alternatively, the overhead stream may be passed to a further reactor in series and treated as described supra.
6 The operating condition such as temperature and gas flow rates 7 would be as described hereinabove.
g The following experimental results are included to describe the operability of the instant invention.
11 A one litre autoclave was fitted with a mechanical 12 stirrer and baffles, a dip tube for sparging N2 or H2 into the 13 reacting liquid, an outlet permitting continuous flow of product 14 gas, and cold trap condensers to remove volatile products from the gas stream before collecting the latter in a sample bag for 16 analysis. Experiments were conducted in the above described 17 reactor at 430C for 105 minutes under 550 RPa pressure both 18 without continuou4 gas flow and with continuous gas flowing in 19 and out of the reactor. In each experiment, 110 g of Athabasca vacuum tower bottom (VTB) and 220 g of a diluent were used.
21 Table I presented herebelow gives the detailed reaction 22 conditions and experimental results.
;`"''''"';
.;
~ . :.' ' '`' .~"~'',' '',. ' .
- ~0048~32 2Batch Autoclave Test Results 3 Test No~ 1 2 3 4 5 6 :
4 Diluent TypeA A B B B B
Nitrogen Flow :
6 Rate (l/min)0.01.89 0-0 0.2 1.05 2.16 -~
7 Yield (wt% VTB):
8 C1-C4 15.5 16.510.3 10.0 8.2 13.8 :~
g Cs - 504C- 42.3 38.156.5 56.~ 56.6 51.4 504C+ pitch 11 (coke free)34.638.826.3 27.0 30.0 29.7 12 Coke 7.9 3.8 4.3 3.8 3.4 3.1 13 Condensate 14 recovered from :.
purge gas (wt% VTB) --47.5 -- 3.1 23.9 39.7 17 Simulated Distillation Results of Condensate from Test 5 18% Off Temp C % Off Temp C
19 IBP 34 55 146 .
24 25 98 80 201 :~
111 85 210 ~-27 40 123 95 229 ~-`
28 45 131 FBP 262 ~ .
29 50 139 -- --- ~
11 ' ' ';
~-:'''' "-' ` ' . - ,.. .
, .', .' 20048~3~
Table III herebelow provides typical diluent and ~-2 Athabasca Vacuum Tower bottom compositions.
4 Diluent Type A B ~ -Weight %~
6 Paraffins 13.0 16.4 ~;
7 Uncondensed 8 Cycloparaffins 7.2 6.3 g Condensed ~ .
Dicycloparaffins 5.2 13.1 ~ :
11 Alkylbenzenes 18.1 15.3 12 Benzocycloparaffins32.3 37.5 :~
13 Benzodicycloparaffin~ 4.8 3.8 ~ ~-14 Naphthalenes 15.9 6.1 . ... -.
Athabasca Vacuum Tower Bottoms (Wt%) ~ -16 C - 81.76 H - 9.51 S -- 6.23 N - 0.78 - -17 API @ 16C : 2.43 ~ ;
"~
-. :'' 18 The large differences in product yield between tests 19 conducted with the respective diluents A and B are due to the high concentration of substituted naphthalenes and tetralins in 21 the latter.
22 In addition to the above described tests, continuous ~ ~
23 flow experiments were performed in a single stirred tank reactor .`
24 and two stirred tank reactors connected in series. A mixture of .:
VTB, diluent, additive (optionally used), and preheated hydrogen ;.
26 were fed into a preheater coil immersed in a heated fluidized l~, '''~' .
20048~
1 sandbath. Hydrocracking reactions were conducted over the 2 appropriate r~sidence time, measured as Liquid Hourly Space ~:
3 Velocity, in the reactor arrangement described in Table IV
4 herebelow. A hot separator split the product into a gas plus volatile stream and a liquid slurry stream.
~`'`
','' ."'''~
~: :
;:0048~32 2 Continuous Bench Unit Test Results ;::
3 Diluent Type : A ~
4 Pressure : 10 NPa ~;
' ',.~, S Test No. 7 8 9 10 6 Reactor 1 (1) 1.2 1.2 1.2 0.5 7 Reactor 2 (1) -- -- -- 0.7 ;~;
8 Reaction Temperature 9 (C) 440 440 460 440 Liquid Hourly ~-11 Space Velocity (hr~l) 1 0.5 0.5 12 Hydrogen Flow rate 13 (slpm) 12 12 12 16 14 VTB Concentration in feed (Wt%~63.10 65.0245.5347.3 16 Conversion (Wt% VTB
17 to coke and 504C-)69.0 78.398.1 79.8 ', 18 Yield, Wt% VTB
19 C1 - C4 5.2 7.9 17.3 7.2 C5 - 200C 16.3 19.1 31.4 15.1 21 200C - 360C 23.4 31.1 43.1 24. 2 : ~:
22 360C - 504C 20. 3 18.4 9. 3 24.8 ~ -~
23 Coke 4.6 3.1 0.1 4.6 24 504C+ Pitch (coke free) 31. 0 21.7 1.9 24.8 ,~
26 Total distillate 27 Cs - 504C 60.0 68.683. 8 64.7 28 It is to be noted that in the test run 9 it was 29 interrupted shortly after the yield period. Plugging occurred downstream of the reactor. Nevertheless, the reactor remained 31 clean in spite of shutdown.
~:.~ ''.'', ' ' `.~ :, ','.',.
, ~
.. . .
-: '~
200~8~
2Table V herebelow shows data with an additive.
TABLE v --~
3 Run No. 11 12 4 Reactor l (1) 1.2 1.2 Reactor 2 (1) 6 Reaction 7 Temperature (C) 450 450 8 Liquid Hourly Space g Velocity (hr~1) 1.05 0.73 (slpm) 8 12 12 VTB Concentration in Feed 13 (wt%) 47.5 47.8 14 Catalyst (Nt~ of Fe based on VTB) 0.5 0.5 ;~
16 Conversion (Wt% VTB
17 to coke and 504C-) 67.8 82.1 18 Yield (Wt% VTB) 19 Cl-C4 5.2 9.4 Cs-200C 19.1 22.9 21 200-350C l9.9 27.6 22 350-504C 21.9 21.1 23 Coke 2.4 2.8 24 504C+ Pitch Coke free 32.2 17.9 ..
26 Total distillate 27 Cs-504C 60.9 71.6 . -~ . .
28 Note that if H2 flow wa~ not increased in the test run 12, one 29 would expect that a 14% increase in conversion should be .::
accompanied by much higher coke yield than the 0.4~ recorded.
:,. ~.
......
'''''~''`
' . ~
- :
Claims (8)
1. A process for hydroconverting a heavy hydrocarbon feedstock which comprises:
- admixing a diluent with the feedstock, and preheating the reaction mixture to a temperature of at least about 350°C; and - contacting the reaction mixture with hydrogen gas in a reactor, said hydrogen gas having a hydrogen concentration greater than 80% and having a flow rate of between about 2,000 - 18,000 scf/BBL, for a liquid hourly space velocity of between about 0.2 to about 2.0, whereby the hydrogen gas flow rate is adapted to substantially immediately strip off the generated lighter components, at a reaction temperature in the range of between about 420°C to about 480°C to thereby selectively hydroconvert said heavy hydrocarbon feedstock and minimize carbonaceous precipitation therein.
- admixing a diluent with the feedstock, and preheating the reaction mixture to a temperature of at least about 350°C; and - contacting the reaction mixture with hydrogen gas in a reactor, said hydrogen gas having a hydrogen concentration greater than 80% and having a flow rate of between about 2,000 - 18,000 scf/BBL, for a liquid hourly space velocity of between about 0.2 to about 2.0, whereby the hydrogen gas flow rate is adapted to substantially immediately strip off the generated lighter components, at a reaction temperature in the range of between about 420°C to about 480°C to thereby selectively hydroconvert said heavy hydrocarbon feedstock and minimize carbonaceous precipitation therein.
2. A process for hydrocracking a heavy hydrocarbon feedstock which comprises:
- admixing a diluent with the feedstock and preheating the reaction mixture to a temperature of at least about 350°C; and - contacting the reaction mixture with hydrogen gas in a reactor operating in a mode approximating a continuous stirred tank reactor, said hydrogen gas having a hydrogen concentration greater than 80%
and having a flow rate of between about 3,000 to about 18,000 scf/BBL for a liquid hourly space velocity of between about 0.2 to about 2.0 at a reaction temperature in the range of between about 420°C to 480°C to thereby selectively hydrocrack said heavy hydrocarbon feedstock and minimize carbonaceous precipitation therein.
- admixing a diluent with the feedstock and preheating the reaction mixture to a temperature of at least about 350°C; and - contacting the reaction mixture with hydrogen gas in a reactor operating in a mode approximating a continuous stirred tank reactor, said hydrogen gas having a hydrogen concentration greater than 80%
and having a flow rate of between about 3,000 to about 18,000 scf/BBL for a liquid hourly space velocity of between about 0.2 to about 2.0 at a reaction temperature in the range of between about 420°C to 480°C to thereby selectively hydrocrack said heavy hydrocarbon feedstock and minimize carbonaceous precipitation therein.
3. A process for hydrotreating a heavy hydrocarbon feedstock which comprises:
- admixing a diluent with the feedstock and preheating the reaction mixture to a temperature of at least about 350°C; and - contacting the reaction mixture with hydrogen gas in a reactor, said hydrogen gas having a hydrogen concentration greater than 70% and having a flow rate of between about 2,000 scf/BBL to about 5,000 scf/BBL for a liquid hourly space velocity of between about 0.2 to about 2.0 at a reaction temperature in the range of between about 420°C to about 455°C to thereby selectively hydrotreat said heavy hydrocarbon feedstock and minimize carbonaceous precipitation therein.
- admixing a diluent with the feedstock and preheating the reaction mixture to a temperature of at least about 350°C; and - contacting the reaction mixture with hydrogen gas in a reactor, said hydrogen gas having a hydrogen concentration greater than 70% and having a flow rate of between about 2,000 scf/BBL to about 5,000 scf/BBL for a liquid hourly space velocity of between about 0.2 to about 2.0 at a reaction temperature in the range of between about 420°C to about 455°C to thereby selectively hydrotreat said heavy hydrocarbon feedstock and minimize carbonaceous precipitation therein.
4. The process as set forth in claims 1, 2 or 3 wherein said diluent is added in the ratio of between about 10:1 to about 1:1 feedstock to diluent and wherein said diluent comprises a product stream generated by cracking heavy oils said diluent having a boiling point in the range of between about 200°C to about 450°C.
5. A process for hydroconverting a heavy hydrocarbon feedstock which comprises:
- admixing a diluent with the feedstock and preheating the reaction mixture to a temperature of at least about 350°C;
- contacting the reaction mixture with hydrogen gas in a first reactor, said hydrogen gas having a hydrogen concentration greater than 80% and having a flow rate of between 2,000 scf/BBL to about 18,000 scf/BBL for a liquid hourly space velocity of about 0.2 to about 2.0 at a reaction temperature in the range of between about 420°C to about 480°C;
- withdrawing the liquid stream from the reactor;
- admixing said liquid stream with hydrogen and a diluent and introducing the mixture to a second reactor;
- contacting the reaction mixture in a second reactor with hydrogen having a flow rate of between about 2,000 scf/BBL to about 18,000 scf/BBL at a liquid hourly space velocity between 0.2 to 2.0 and at a reaction temperature in the range of between about 420°C to about 480°C to thereby selectively hydroconvert said heavy hydrocarbon feedstock minimizing carbonaceous precipitation therein.
- admixing a diluent with the feedstock and preheating the reaction mixture to a temperature of at least about 350°C;
- contacting the reaction mixture with hydrogen gas in a first reactor, said hydrogen gas having a hydrogen concentration greater than 80% and having a flow rate of between 2,000 scf/BBL to about 18,000 scf/BBL for a liquid hourly space velocity of about 0.2 to about 2.0 at a reaction temperature in the range of between about 420°C to about 480°C;
- withdrawing the liquid stream from the reactor;
- admixing said liquid stream with hydrogen and a diluent and introducing the mixture to a second reactor;
- contacting the reaction mixture in a second reactor with hydrogen having a flow rate of between about 2,000 scf/BBL to about 18,000 scf/BBL at a liquid hourly space velocity between 0.2 to 2.0 and at a reaction temperature in the range of between about 420°C to about 480°C to thereby selectively hydroconvert said heavy hydrocarbon feedstock minimizing carbonaceous precipitation therein.
6. The process as set forth in claims 1, 4 or 5 wherein said hydrogen partial pressure ranges from between about
7 MPa to about 15 MPa.
7. The process as set forth in claim 6 which further comprises admixing an iron, nickel, molybdenum or cobalt carbonyl or iron sulphate with said reaction mixture.
7. The process as set forth in claim 6 which further comprises admixing an iron, nickel, molybdenum or cobalt carbonyl or iron sulphate with said reaction mixture.
8. The process as set forth in claims 2, 3 or 4 wherein said hydrogen gas is introduced into the reactor in countercurrent flow with said feedstock.
Priority Applications (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA 2004882 CA2004882A1 (en) | 1989-12-07 | 1989-12-07 | Process for reducing coke formation during hydroconversion of heavy hydrocarbons |
Applications Claiming Priority (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA 2004882 CA2004882A1 (en) | 1989-12-07 | 1989-12-07 | Process for reducing coke formation during hydroconversion of heavy hydrocarbons |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA2004882A1 true CA2004882A1 (en) | 1991-06-07 |
Family
ID=4143744
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA 2004882 Abandoned CA2004882A1 (en) | 1989-12-07 | 1989-12-07 | Process for reducing coke formation during hydroconversion of heavy hydrocarbons |
Country Status (1)
| Country | Link |
|---|---|
| CA (1) | CA2004882A1 (en) |
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- 1989-12-07 CA CA 2004882 patent/CA2004882A1/en not_active Abandoned
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